Benzene-naphtha reforming process



Apri 25, w67 w. F. JQHNSTON, JR Ew BENZENE-NAPHTHA REFORMING PROCESSFiled May 14, 1964 1 Sm t United States Patent O 3,316,316BENZENE-NAPHTHA REFRMIN G PRUCESS Walker F. Johnston, Jr., Flossmoor,Ill., assignor to Standard Gil Company, Chicago, Ill., a corporation ofIndiana Filed May 14, 1964, Ser. No. 367,334 9 Claims. (Cl. 260-668) Myinvention relates to the production of ultra high octane number blendingstocks from naphtha feed stocks and oleiinic benzene-containing materialboiling in the gasoline boiling range. More particularly my inventionrelates to a combined process of catalytic reforming, hydrogenation ofthe olefinic benzene-containing material and introduction of thishydrogenated material into the reforming process at a selected point toproduce a reformate product from which high purity benzene and ultrahigh octane gasoline blending components may be easily recovered in highyield.

The primary object of the invention is to provide means for economicallyprocessing olefinic benzene-containing materials, such as steam-crackednaphthas and dripo` lene, and naphtha to produce in high yield a highoctane reformate suitable for use as a gasoline blending component.Another object is a combination process for the recovery of highquality, low sulfur benzene from dripolene without destroying anappreciable amount of the benzene contained therein and without resortto elaborate and costly treating and purification processes. The presentinvention provides means whereby substantial savings in catalystrequirements for producing high quality benzene and 10U-{- Researchoctane number gasoline blending stocks can be achieved and also providesmeans for maximizing the recoverable liquid yield of benzene and highoctane product.

It is well known that the high temperature pyrolysis of gaseoushydrocarbons to prepare ethylene results in the by-production of anormally liquid mixture of hydrocarbons through reactions such aspolymerization, alkylation, aromatization, dehydrogenation, and thelike. The mixture is commonly termed dripolene, and while it containsvirtually all classes of hydrocarbons it predominates in oleiinic andaromatic hydrocarbons, mainly benzene. As the demand for ethylene forthe production of polyethylene and plastics and other petrochemicalsrises, increasingly large supplies of dripolene are becoming available.Because of its benzene content, dripolene represents an exceedinglyvaluable material, and one which is steadily becoming more obtainable.

Thus far lthree large-volume usages have developed for dripolene. It isblended into motor gasolines, where the high octane numbers of itsaromatic and olenic components render dripolene a desirable blendingstock. Dripolene may be fed to aromatics extraction units for therecovery `of benzene. Finally, exceedingly valuable resins have beenmade by thermal or catalytic polymerization of high-boiling dripolenefractions.

Dripolene however contains cylic di-oleiins and other reactive oleliniccompounds which give rise to problems which have severely restricted thequantity of dripolene blended into motor fuels or fed to aromaticsextraction units. Cyclic di-olefins tend to form gum-like polymers inthe presence of air or upon heating, and for this reason only relativelysmall amounts, usually less than about 0.5 volume percent, of dripolenecan be blended into premium motor fuels. And in aromatics extractionunits, it is r'ce found that cyclic di-olefins tend to concentrate inthe aromatic extract, thereby complicating the preparation of purearomatic compounds, such as benzene.

A catalytic hydrogenation process has previously been discovered wherebydripolene may be hydrogenated successfully lto remove these di-olens andother gum-formers essentially completely by converting them to saturatedor less olefinic compounds, This process is `described in U.S. PatentNumber 2,953,612. However, the benzene fraction, boiling in the range ofabout -190 F., of hydrogenated dripolene still contains an appreciableamount of olenic and sulfur compounds. Many commercial processesutilizing benzene as a starting material require high purity benzenecontaining essentially no oleiinic or sulfur impurities, therefore thebenzene recovered from hydrogenated dripolene must be further purifiedbefore it is marketed. Also the F. plus fraction of hydrogenateddripolene has a clear Research octane number of about 98.5 which, bymodern standards, is relatively low. Since current premium motorgasolines are being marketed having octane numbers above 100, it isdesirable that the research octane number of the hydrogenerateddripolene components other than benzene be increased in order to enhancetheir usefulness as premium gasoline blending components.

Previous attempts to recover benzene and high octane reformate fromdripolene by including even small concentrations of dripolene with thenormal naphtha feed to catalytic reformers have been unsuccessfulbecause of fouling of pre-heat exchangers and furnaces, as well as thecatalyst beds themselves, due to polymerization and condensation of theolenic materials contained in the dripolene. These fouling problems wereovercome by hydrogenating the -dripolene prior to including it in thehydroformer feed. While this mode of operation is generally successful,much of the benzene present in the hydrogenated dripolene, i.e. up toabout 30%, is lost in the reforming process by conversion to othercompounds.

In the catalytic reforming of naphthas several reactions occur,principally dehydrogenation, isomerization, dehydrocyclization andcracking. The dehydrogenation reaction is the easiest and fastest of thereactions, therefore it takes place in the lirst portion of thereforming catalyst. An example of this reaction is dehydrogenation ofcyclohexane to produce benzene and hydrogen. Since the dehydrogenationreaction is highly endothermic, and some of the other reactions are alsomildly endothermic, catalytic reforming processes normally employ aplurality of catalyst beds or zones, termed reactors herein, withprovisions for reheating the process stream between the beds or zones.In such multi-reactor catalytic reformers the dehydrogenation reactionis essentially complete in the first one or two reactors. My inventioncontemplates treating the -rst reactor or reactors to be contacted bythe naphtha feed as the dehydrogenation reactors which I will designateas the dehydrogenation stage or simply as the first stage. In thesubsequent stage, which comprises the remaining reactors or zones, theisomerization, dehydrocyclization and cracking reactions predominatesince the dehydrogenation reaction is essentially completed in the firststage.

I have found that when hydrogenated dripolene is introduced into theprocess stream of a naphtha reformer subsequent to the dehydrogenationstage no loss of benzene occurs. The components of hydrogenateddripolene, other than benzene, are up-graded in research octane ratingand he sulfur and olefin impurities present in the benzene fracion ofthe hydrogenated dripolene are simultaneously renoved so that highpurity benzene may be easily recovered rom the reformate product withessentially no loss of the tenzene present in the hydrogenateddripolene. Thus, the )rocess of my invention results in the productionof high luantities of high quality benzene and high octane reormate fromnaphtha and hydrogenated dripolene with ess processing equipment than inany other known process.

I prefer to limit the amount of hydrogenated dripolene which isintroduced into the reformer below a 1:1 volume .'atio of hydrogenateddripolene to naphtha charge. A :onvenient amount of dripolene overheadcharge is that amount which will result in the net hydrogen produced inthe reformer being consumed in the hydrogenator so that the combinationprocess is operated in hydrogen balance. This results in maximumutilization of the chemically combined hydrogen present in the naphthafeed. lt is desirable to desulfurize the naphtha feed to the reformer inorder to reduce the load on the H28 removal facilities and minimizecorrosion of process equipment.

Although the amount of hydrogen consumed in the hydrogenation of -thedripolene overhead varies with the properties of the particulardripolene used, the hydrogen consumption is normally in the range of 500to 750 standard cubic feet per 4barrel of dripolene. Likewise the yieldof hydrogen from the reformer will vary depending upon the properties ofthe particular naphtha feed. Reforming of South Texas heavy naphthatypically yield-s 1,000 standard cubic feet (s.c.f.) of hydrogen/ bbl.of feed. In this situation, a dripolene overhead/ naptha volume ratio ofabout 1.3:1 to 21:1 results in operation of the process of the inventionin hydrogen balance. Of course, if Asome of the reformer hydrogen isused to hydrodesulfurize the naphltha feed, this ratio is loweredcorrespondingly.

Example To illustrate my invention and the advantages thereof, l havechosen for dripolene hydrogenation the process of U.S. Patent 2,953,612and a 5-reactor catalytic reformer utilizing platinum-alumina-halogencatalyst. The reformer catalyst has a platinum content in the range of0.1 to 1 weight percent and is disposed in 5 reactors with furnacesprovided prior to each reactor for heating the process stream. Thecatalytic reforming process operates at a pressure in the range of 50 to500 p.s.i.g., average reactor temperatures in the range of 750 to 50 F.,and a weight hourly space velocity (Wo/hn/Wc) in the range of 0.5 to 5.Hydrogen-containing gas is recycled in an amount to provide about 2500to 6000 standard cubic feet of hydrogen per barrel of naphtha feed.

Three modes of operation are compared:

(l) Separately reforming naphtha and recovery of ben zene fromhydrogenated dripolene,

(2) Reforming naphtha containing 15 vol. percent of hydrogenateddripolene,

(3) And reforming naphtha with introduction of hydrogenated dripoleneinto the reforming process stream subsequent to the first stage, whichin this 5-reactor system is subsequent -to the second reforming reactor.

The quantities of naphtha and hydrogenated dripolene utilized in each ofthe three modes of operation and the reforming severity in each case arethe same to permit ready comparison. Benzene recovery in each case iseffected by separating from the reformate a narrow-boiling fraction ofbenzene concentrate and recovering the benzene therefrom by extractivedistillation with phenol. The quantity of high quality benzene and theamount and octane of the reformate produced by each of the three modesof operation are compared in Table I. These data show that operationaccording to the present invention, Mode 3, produces more high octanereformate than either Mode 1 or Mode 2. Also the quality of the benzeneproduced in Mode 3 is higher than that produced 4 according to Mode 1and more of it is produced than by Mode 2. Although these differencesmight at rst glance appear small, these are valuable products which areproduced in great quantities, therefore the monetary advantage ofoperating according to the present invention is indeed substantial.

i Hydrogenated Dripolene overhead, not. reformed.

2 lyllixtlre of 85 bbls. paphtlia and 15 bbls. hydrogenated dripoleneover rea 3 85 bbls. naphtha [ed to reformer, 15 bbls, hydrogenateddripolene overhead introduced to reformer subsequent the dehydrogenatioustage.

4 C-lreformate, less benzene. HDO not reformed in Mode 1.

According to the invention, the naphtha feed is charged to a reformingZone in which at least three reactors connected in series are provided.1t is preferred that the naphtha feed be desulfurized. Anydesulfurization process known to the art may be employed. The reactorsare equipped in the usual manner for inter-heating between reactors inorder to compensate for endothermic temperature drop and to introduceadditional heat into the later stages of the reaction. The reactorscontain a platinum type reforming catalyst in the form of pelleted,pilled, extruded or beaded particles. In the reforming zone, separatedehydrogenation and isomeriZation-dehydrocyclization-cracking stages arereadily identified by the relatively large endothermic temperature dropacross the catalyst beds of the first one or two reactors encounteringthe naptha charge.

The efuent from the reforming zone is separated conventionally into arecycle hydrogen gas fraction and a liquid reformate fraction. Thelatter, if separate benzene recovery is desired, is further fractionatedso as to separate light hydrocarbons boiling in the C6 range and torecover a C74- fraction having an ultra high research octane rating.With a commercially available platinumalumina catalyst as used inUl-traforming, for example, under `reforming conditions including apressure in the range of about 15 0` to 400 p.s.i.g., the severity levelshould be sufficient to produce at least -a 95 research octane C5|reformate from a Mid-Continent heavy naphtha charge, and moreadvantageously, such as to produce a `100-1- research octane C54-reformate.

Dripolene normally boils in the range of about 100 to 400 F. asdetermined by ASTM distillation, and may contain appreciable amounts oflight, normally gaseous hydrocarbons. My invention is particularlyconcerned with the fraction of dripolene boiling in the range of about100-375 F., although it is not essential that the dripolene boilentirely over this range or that all of the dripolene fraction boilwithin the range. My dripolene charging stock is obtained as an overheador heartcut fraction in the distillation of total dripolene to obtainabout to of the charge fraction, which I term dripolene overhead, whilethe bottoms 4may be used to prepare resins by processes well known -tothe art.

Dripolene overhead is hydrogenated in the presence of a platinum-aluminacatalyst at elevated tem-peratures and pressures and in the presence ofhydrogen gas to selectively hydrogenate monoolens and diolens withoutsimultaneously hydrogenating the aromatic compounds to naphthenes.

The dripolene fraction and a hydrogen-containing gas stream initiallycontact the platinum-alumina hydrogenation catalyst at conditions oftemperature and pressure `such that substantially -all (ie. about 80% ormore, and preferably at least 90%) of the dripolene remains in theliquid phase, while maintaining a critically low hydrogen sulfideconcentration in the hydrogen-containing gas stre-am.

The a-mount of hydrogen sulfide that may be present in the process isquite critical, and accordingly it is essential to provide to thereaction zone a hydrogen-containing gas stream containing less than 12grains of hydrogen sulfide per 100 standard cubic feet of gas. If forexample the level of sulfur in the hydrogen-containing gas exceeds 12grains, both the catalyst activity `and the catalyst life diminishrapidly, and if the level increases to as much as 57 grains per 100s.c.f., the catalyst becornes completely deactivated in a matter ofminutes. Fortunately however, the effect of either hydrogen sulfide gasor merca-ptan sulfur in the dripolene feed on the hydrogenation process-appears to be temporary with respect to its effect on product qualityalthough periods of high sulfur do materially increase the amount ofcoke deposition. It is convenient to use the hydrogen-rich gas producedby the catalytic reformer as the source of hydrogen for thehydrogenation process.

Suitable platinum-alumina hydrogenation catalysts are conveniently thosecatalysts which have been found eminently suitable for use in naphthareforming processes. Generally, these catalysts contain from about 0.01to about by weight of platinum and may optionally from -about `0.05 toabout 3% by weight of a halogen, preferably chlorine and/or fluorine, ona high surface area alumina support such as the alumina described inHeard Reissue Patent Number 212,196. The catalysts may be in the formIof pills, pellets, extrudates, spheres or the like, and conventionallyhave a size between about y1@ to 1A in maximum dimension. A particularlysuitable catalyst is one which has been partially deactivated bycontinued use in a naphtha reforming process, since it appears thatcatalysts previously used for reforming are more stable and have lesstendency to hydrogenate aromatic compounds than fresh platinum-aluminacatalysts. Furthermore, used catalysts exhibit less tendency to causewasteful 'hydrocracking of hydrocarbons and thus result in higher yieldsof recoverable liquid product.

The conditions of pressure, temperature, liquid hourly space velocityand hydrogen-containing gas rate which are employed are interrelatedsuch that the commingled feedstock and hydrogen-containing gas, as itinitially contacts the catalyst bed, consists of la gas phase and aliquid phase wherein the liquid phase comprises substantially all of thecharge stock. It is desired that at least 80 mol percent, and preferablyat least 90 mol percent of the dripolene charge contact the catalyst `asa liquid. Pressures within the range of 100 to 1000 p.s.i.g. aredesired, with pressures fr-om 300 to 500 p.s.i.g. preferred from acommercial standpoint as this latter range favors conditions at whichthe hydrogenation reaction occurs rapidly. `Within the broad presurerange the bed inlet tem-perature may be between 50 and 200 F., mostdesirably between 100 and 150 F., typically 115 F. With most dripolenestocks the temperature rise through an adiabatic bed, for completeoleflnic saturation, is on the order of 350-450 F. and provides anaverage reactor temperature of about 280-340 F. This average temperaturemay be increased by providing more catalyst or may be decreased byincreasing the proportion of hydrogen-containing gas to charge stock.With respect to the hydrogen-containing gas, it is desirably employed ina proportion of 500 to 10,000 standard cubic feet per barrel`(s.c.f./b.) of charge stock, preferably from 1000 to 4000 s.c.f./b.,e.g. 1500 s.c.f./b. This gas preferably is cornposed of at least 70%hydrogen as derived from the naphtha hydroforming operation. Althoughthe experi- .mentally observed consumption of hydrogen usually variesbetween about 5 00 and 750 s.c.f./b., it is preferred to maintain asubstantially large amount in the reaction zone.` This may beaccomplished economically be relcycling the excess hydrogen. Thehydrogen-containing gas, if recycled, must be chemically treated tomaintain the critically low hydrogen sulfide level therein.

The effluent from the hydrogenation reactor is introduced into thereforming process stream subsequent to the first, or hydrogenation,stage. The entire eflluent stream may be so introduced, or ahydrogen-rich gas may 'be recovered from the effluent for recycle to thehydrogenation reactor, and only the hydrocarbon portion of the effluentintroduced into the reformer. A convenient way to recover :ahydrogen-rich gas for recycle is to cool the eflluent to condense thenormally liquid hydrocarbons therein and pass the cooled stream to agas-liquid separa- Itor from which hydrogen-rich gas and liquidhydrogenated dripolene may be withdrawn.

Simultaneous reforming o-f the effluent from the reformer hydrogenationstage and the introduced hydrogenated dripolene is then effected in thesubsequent stage of the reformer. Then, if desired, the C6 reformate isseparated by fractionation into a fraction boiling in the range of aboutto 190 F. and a heavy reformate fraction having an initial boiling pointin the range of about to 230 F. which contains at least about 85 yvolumepercent aromatics. The latter constitutes a product stream having aclear research octane number significantly in excess of 100. The 140 to190 F. cutis subjected to solvent extraction or extractive distillationto recover high purity benzene. Alternatively, the total reformate maybe used as a high octane gasoline blending component having an ususuallyhigh front-end octane because of the increased benzene content.

The invention will be further described by reference to the accompanyingdrawing which is a flow plan of a preferred embodiment of the inventionin simplified diagrammatic form.

The feed, constituting a 200 to 400 F. mixture of South Texas naphthas,is charged to the system through line 10. The feed is preheated in firedheater 11 and is mixed in line 12 with recycle hydrogen gas from line13. The mixture is charged to reactor 14, which is the first of a trainof ve serially connected reactors, each of which contains a bed ofplatinum-alumina catalyst in pellet form. The reaction mixture is flowedfrom reactor 14 via line 14a to interheater 16 and from thence viaconnection 17a to reactor 17. The effluent from reactor 17 is passed bymeans of l-ine 18, into which hydrogenated dripolene from line 66 ischarged, through interheater 19 and from thence by means of connection20a to reactor 20. From reactor 20, the reaction mixture is passed bymeans of line 21, interheater 22 and connection 23 to reactor'24. Fromreactor 24, the reaction mixture is passed by means of line 25,interheater 26 and connection 27 to the last reactor 28. Hydrogenateddripolene may also be introduced into the reformer process stream viavalved line 67 :and/or 68.

The eflluent from the final reactor 282 is flowed through line `29 andcooler 30 to high pressure gas separator 31. In separator 31, a recyclegas rich in hydrogen is recovered by line 32 for recompression andrecycle through line 33, heater 34 and line 13. Hydrogen-rich make gasis Withdrawn from line 32 via line 3S for use in hydrogenating thedripolene. Excess make gas may be vented Ifrom the system through valvedconnection 52.

Turning noW to the dripolene hydrogenation, dripolene liquid iswithdrawn from external storage tanks and conducted through line 36 tofractionator 37 which is provided with corrosion resistant distillationtrays or perforated pans, wherein an overhead dripolene charge stockfraction comprising about 80% of the total dripolene is separated `bydistillation from about 20% of high boiling bottoms, which latter issent via line 38 to the resins plant, not shown. The total dripolene fedto fractionator 37 has an analysis approximating the typical dripolenedescribed 7 reviously. The 80% fractionator 37 overhead which is vkenthrough line 39 has the following compositi-on:

TABLE I1 'harge analysis:

Gravity, A.P.I 32.4 R.V.P.,p.s.i.a 6.18 LSTM distillation, F.:

I.B.P. 134 163 179 50% 1189 70% `202 90% 280 F.B.P. 356

Light hydrocarbons anlysis )omponentz C3 liquid vol. percent 0.1 fc4:dO-.... 1C4,= do 0.2 210,: do 0.2 nC4= do 0.2 C., diolen do 1.6 C5diolen do 6.6 C5 monoolen do 3.1 C5 paraffin do 0.2 Cs-ido 87.5 Benzenedo 53 C64-gravity, A.P.I 28 0 The 80% dripolene charge fraction contains70 parts per million sulfur and 29 parts per million organic chlorides,and has a bromine number of 48 (indicative of total olens) and a maleicanhydride value (MAV, representing conjugated diolefins) of 47 Ing/g.The bottoms withdrawn through line 38 has an ASTM distillation boilingrange between about 200` and 420 F., preferably between about 230 and375 F.

The dripolene charge is conducted through line 39, cooler 40, and line49 to a charge pump, not shown, which may be a multistage centrifugalpump adapted to pump the dripolene charge to the reactor systemoperating at a pressure of 325 pounds per square inch gage. The cooler40V outlet temperature is about 80 F. The charge stock from the pump issent through line 41 to junction 42, where it is met by a stream ofrecycle hydrogen-containing gas from line 43 in the amount of 1350standard cubic feet of total hydrogen-containing gas per barrel ofcharge. The gas has a composition of approXimately 80% hydrogen, withthe balance consisting primarily `of methane, ethane, and some propaneand propylene, together with less than the critical limit of 12 grainsof H28 per 100 `standard cubic feet of total gas. It is highly preferredthat this gas contain, if possible, less than 3 grains per 100 cubicfeet of HZS. The temperature of the commingled liquid and gas stream is115 F., and at this temperature the commingled stream passes via line 44into reactor 45 shown symbolically as a single bed or chamber, althoughit may comprise a plurality of serially lor parallel-connected reactionchambers. At these operating conditions, 94 mol percent of the dripoleneis in the liquid phase when the commingled stream initially contacts thecatalyst.

The reaction zone 45 operates essentially adiabatically, that is thecommingled dripolene charge and hydrogencontaining gas stream arepermitted to increase in temperature by the eXothermic heat of monoolenand diolefin hydrogenation on passage through the catalyst bed. Thecatalyst employed is spent Ultraforming catalyst obtained after morethan one years use in a regenerative naphtha reforming unit and has anactivity for reforming of substantially less than that of freshUltraforming catalyst, but is very nearly as active for hydrogenation asis fresh catalyst. The catalyst in chamber 45 is in the form of pelletshaving an average length and diameter approximating 1/8 and is disposedso as to permit downllow passage of the commingled stream. A Weighthourly space velocity of 2 is employed. In passage through the reactionzone the dripolene plus hydrogen stream temperature is increased to 625F., which provides an average reaction temperature of 370 F. In thiszone, 625 standard cubic feet per barrel of hydrogen is consumed byolefin hydrogenation, a quantity which compares closely with thetheoretical hydrogen consumption based on the observed experimental heatof reaction, 280 B.t.u./ lb. The quantity of catalyst in reaction zone45 is that which provides a weight hourly space velocity of 2.0, i.e.2.0 pounds of dripolene charged per hour for each pound of catalyst inZone 45. The hydrogenated stream leaving chamber 45 passes through line46, valved line 47, to cooler 48, and then through line 49 to gas-liquidseparator 50.

The hydrogenated product stream comprising hydrogenated dripolene invapor form together with excess hydrogen-containing gas is cooled in thecooler 48 wherein the hydrogenated dripolene condenses as a liquid whichis sent, along with the non-condensible hydrogen-containing gas, to thegas-liquid separator 50. At the gas-liquid separator 50, thehydrogen-containing gas is separated and withdrawn through line 51 andconducted to amine scrubber 53, where a descending stream ofdiethanolamine or other agent, from line 60, effective to absorb HZS isemployed to remove hydrogen sulfide gas formed by the destructivehydrogenation of sulfur compounds in the dripolene charge or in thenaphtha reformer feed and carried in the reformer make gas stream feedto the hydrogenation system via lines 35 and 54. The amine is withdrawnvia line 55, heated in a stripper, not shown, for the purpose ofreleasing absorbed H28 and recycled to the amine scrubber 53.

Depending upon the hydrogen sulfide concentration of the reformerhydrogen-containing gas, it may be added at either valved line 54, orvalved line 43. Briefly, if the reformer gas is relatively low inhydrogen sulfide, it may be added to the system through valved line 43.The cornposition -of reformer gas varies with the operation of thereformer and may range for example from 70-95% hydrogen, the balancebeing saturated light hydrocarbons such as methane, ethane and propane.If this gas is of a purity below about it may be desirable 'to vent aportion of the gas from gas-liquid separator 50 through valved vent line57 so as to prevent a build-up of noncondensible methane, ethane andpropane within the recycle gas system.

Where large quantities of reformer gas are available, the presentrecycle gas system may be eliminated in favor of a once-through hydrogenow. In this case, the entire hydrogenation reactor efuent may be passeddirectly via line 46, and valved line 69 into line 65 for introductioninto the reformer process stream.

After treatment in amine scrubber 53, the essentiallyhydrogen-sulfide-free hydrogen-containing gas, .e. containing less than12 grains of HZS per 100 standard cubic feet, is conducted via line 58to water -scrubber 59 where a descending stream of water from line 61scrubs entrained or vaporized amine from the gas. The rich water streamis withdrawn through line 62 and is concentrated for amine recovery in adistillation column, not shown. If desired, water vapor removalfacilities such as a glycol scrubbing tower or a silica gel or aluminadrier may follow water scrubber 59 in line 63.

The treated gas passes from water scrubber 59 through lines 63 and 43back to the juncture 42 with dripolene charge line 41 and thence toreaction Zone 45.

Returning now to receiver 50, hydrogenated dripolene as a liquidcondensate passes through lines 64, 65 and any one or more of valvedlines 66, 67, and 68 into the reformer process stream in one or more oflines 18, 21 and 25.

The average pressure on the reforming system is 300 p.s.i.g., the spacevelocity is 1.0 weight of fresh feed per hour per weight of catalyst,and the recycle rate is 5,000 s.c.f. of hydrogen per barrel of feed. Thefeed is preheated to a temperature of 900 F. in heater 11, and therecycle hydrogen is heated to a temperature of 1025 F. in heater 34,providing a reactor inlet temperature of 940 F. The outlet temperatureof reactor 14 is 800 F. The reaction mixture is reheated in interheater16 to a temperature to provide an inlet temperature to reactor 17 of 940F. The outlet temperature is 880 F. In similar fashion, the feed streamsto the remaining reactors are reheated to obtain inlet temperatures of940 F. The outlet temperatures for reactors 20, 24 and 28 are,respectively, `910", 925 and 930 F.

The effluent from reactor 28 is cooled to obtain a temperature of100-120 F. in gas liquid separator 31 at 280 p.s.i.g. The octane numberof the usual (35+ reformate obtained from the South Texas charge naphthais 100 research octane clear under the above reaction conditions. Thisreformate may be withdrawn via valved line 70 for utilization as a highoctane gasoline blending component. Alternatively, when high puritybenzene is to be recovered, the reformate from the separator 31 is fedto fractionator 71 wherein a heavy reformate having an initial boilingpoint of about 230 F. and having a research octane number of about 108research octane clear, comprising about 54 volume percent of the productis recovered as bottoms from fractionator 71 via line 72. Thebenzene-containing overhead fraction from fractionator 71 is passed vialine 73 to a solvent extraction tower 74 wherein the benzene-containingoverhead fraction is contacted with diethyleneglycol solvent at asolvent to oil ratio of 6:1, 300 F. and 150 p.s.i.g. The rich solventstream containing the benzene is withdrawn from the bottom of theextraction tower 74 and passed via line 75 into solvent stripper 76wherein benzene is distilled overhead from the solvent and withdrawn vialine 77. A lean solvent is withdrawn from the bottom of stripper 76 vialine 78 and recycled to the extraction tower 74 via line 79. Raflinateis withdrawn from the top of extraction tower 74 via line 80. Theraffinate may be withdrawn from the system via valved line 81, oralternatively, the rafnate may be recycled to the reforming process viavalved line 82, line 65 and any one or more of valved lines 66, 67 and68.

In the operation of the invention, three or more reactors can be used inthe reforming system. In the case of three reactors, the first reactorconstitutes the dehydrogenation stage. In the `case of a greater numberof reactors, the first two reactors will usually constitute thedehydrogenation stage. The conditions in the dehydrogenation stageapproximate 750 to 900 F. average -temperature, with a space velocity inthis first stage, based on naphtha feed, of about 3 to 10 WHSV. Thetemperature in the subsequent stage will approximate 900 to l,000 F. Thepressure may be in the range of 100 to 500 p.s.i.g., preferably 150-300, and the hydrogen recycle rate in the range of about 2,000 to10,000 s.c.f. per barrel.

The reforming catalyst may comprise any of the platinum-type reformingcatalysts, preferably on an alumina type base, although other supports,such as deactivated silica-alumina, alumina-titania, and the like, mayhe used. The presence of chlorine or iluorine, in known manner, may bedesired in the reforming zone.

Although the drawing illustrates only the hydrocarbon flow, it will beunderstood that the system should be equipped for catalyst regenerationand/or rejuvenation. In the regeneration step, carbon is burned off thepartially deactivated catalyst with -a dilute oxygen containing gas.Higher oxygen partial pressures and severities are used in rejuvenationof more severely deactivated catalysts. The regeneration may be effectedperiodically in blockedout operations, or it may be effected in themanner of ultraforming by use of a swing reactor as has been describedin the technical literature.

The cut point `in the reformate splitter depends sorne what upon theseverity of reforming, the feed stock and the desired heavy reformateoctane. With C7+ charge naphthas at severity levels producing octaneproduct, the C7 aromatics should be included in the heavy reformate. Aninitial boiling point in the range of about 225 to 250 F. lisrecommended. At lower severities or with more refractory feeds, theinitial of the heavy reformate may be in the range of about 250 to 275F., however, this will require an additional distillation `step toseparate benzene and toluene recovered from the extract.

A variety of extractive agents can he used inthe solvent extraction fortreating the light reformate. Polyhydroxy solvents such asdiethyleneglycol, dipropyleneglycol, triethyleneglycol, or mixturesthereof, `advantageously promoted in selectivity by the addition ofwater, `are particularly suitable. Other useful solvents are describedin U.S. Patent 2,365,517. A newer solvent, butyrolacetone, has certainadvantages for processing reformates. Sulfur dioxide also is feasiblealthough it requires added facilities for refrigerated handling of thesolvent. Usually, the feed and solvent will be contactedcountercurrently in one or more extraction columns of the number oftheoretical extraction stages required to effect the degree ofseparation desired. The conditions of extraction will be determined vbythe nature of the solvent and its selectivity for aromatics at varioustemperature conditions. Usually, selectivity is improved with decreasingtemperatures, and temperatures in the range of say -40 F. to 300 F. ormore may be used, with adjustment of pressure to obtain the desiredphase separation, at solvent to feed ratios advantageously in the rangeof l/ 1 to 25/1. Various methods may be used to separate extract andsolvent, but, in general, distillation is most satisfactory. Traces ofsolvent can he removed from the separated rainate and extract phases bywashing with water or other solvents, or by stripping.

The invention has the advantage of improving reformate yield at anygiven severity yby introducing greater selectivity into the conduct ofthe various reforming reactions and of producing high quality,lowrsulfur benzene. The most difficult reforming reaction, cyclizationof parans, is promoted in rate fby decreasing the concentration ofnaphthenes in the subsequent reaction zone. Thus, injection of thehydrogenated dripolene has the additional advantage of reducing theeffective concentration of naphthenes to a very low level. Also, moreselective handling of the hydrocarbons in the feed is possible sincegreater advantage is taken of the different conditions obtaining througha series of reforming reactors by injecting the hydrogenated dripoleneand, if desired, the paraifnic raffinate to the latter reactors.

I claim:

1. A method of producing high quality benzene from dripolene whichcomprises hydrogenating dripolene, thereafter charging hydrogenateddripolene to a naphtha reforming process comprising dehydrogenation,isomerization, dehydrocyclization and cracking reactions subsequent thedehydrogenation reaction, withdrawing reformate, and recovering highquality benzene therefrom.

2. A method Iof producing benzene and reformate from olerinicbenzene-containing material boiling in the gasoline boiling range andnaphtha which comprises hydrogenating said benzene-containing materialto reduce the olefin and diolefin content thereof, charging said naphthaand hydrogen to the dehydrogenation stage of a reformer, charging saidhydrogenated benzene-containing material and eluent from saiddehydrogenation stage to a subsequent stage of said reformer andrecovering benzene and reformate from the effluent of said. reformer.

3. A method of producing benzene and reformate from oleflnicbenzene-containing material boiling in the gasoline boiling range andnaphtha which comprises hydrogenating said benzene-containing materialto reduce the fouling tendency thereof, charging said naphtha andhydrogen to the dehydrogenation stage of a reformer, charging saidydrogenated benzene-containing material and efuent rom saiddehydrogenation stage to a subsequent stage of aid reformer, recoveringbenzene, reformate and hydroen-rich gas from the eiuent from saidreformer, recycling portion of said recovered hydrogen-rich gas to thedeiydrogenation stage of said reformer to provide said hylrogen, andemploying another portion of said recovered iydrogen-rich gas in saidhydrogenation.

4. An improved method of concurrently reforming iaphtha and hydrogenateddripolene which comprises hylrogenating said dripolene to producehydrogenatcd drip- )lene, charging said naphtha and hydrogen to thefirst stage of a reformer, charging both of said hydrogenated :lripoleneand hydrocarbon eiuent from said rst stage to :he second stage of saidreformer, and recovering reformate from the efuent of said second stage,said reformate having a higher yield and/or octane rating than if saiddripolene were fed to the rst stage of said reformer in the Conventionalmanner.

5. An improved method of concurrently reforming naphtha and hydrogenateddripolene which comprises separately desulfurizing said naphtha toproduce desulfurized naphtha and hydrogenating said dripolene to producehydrogenated dripolene, charging said desulfurized naphtha to thedehydrogenation stage of a multi-stage reformer, charging both of saidhydrogenated dripolene and effluent from said dehydrogenation stage to asubsequent stage of said reformer, and recovering reformate from theeffluent of said subsequent stage.

6. An improved process for producing benzene and reformate from naphthaand dripolene which comprises hydrogenating said dripolene to producehydrogenated dripolene, charging said naphtha to the first stage of amulti-stage reformer, charging said hydrogenated dripolene and effluentfrom said rst stage to the second stage of said reformer, recoveringbenzene from the efiluent of said second stage and recovering at least aportion of the remainder of said second stage efuent as reformate,whereby the yield of benzene is higher than when said hydrogenateddripolene is charged to said reforming rst stage.

7. The method of claim 1 wherein said reforming process employs aplatinum-alumina catalyst.

8. The method of claim 1 wherein said hydrogenating is carried out inthe presence of a platinum-alumina catalyst.

9. The method of claim 1 wherein the hydrogen consumed by saidhydrogenating is produced in said reformer.

References Cited by the Examiner UNITED STATES PATENTS 2,865,837 12/1958Holcomb et al. 208-138 2,953,612 9/1960 Haxton et al 208-144 DELBERT E.GANTZ, Primary Examiner.

A. RIMENS, Assistant Examiner'.

1. A METHOD OF PRODUCING HIGH QUALITY BENZENE FROM DRIPOLENE WHICHCOMPRISES HYDROGENATING DRIPOLENE, THEREAFTER CHARGING HYDROGENATEDDRIPOLENE TO A NAPHTHA REFORMING PROCESS COMPRISING DEHYDROGENATION,ISOMERIZATION, DEHYDROCYCLIZATION AND CRACKING REACTIONS SUBSEQUENT THEDEHYDROGENATION REACTION, WITHDRAWING REFORMATE, AND RECOVERING HIGHQUALITY BENZENE THEREFROM.